Process for fluid catalytic cracking

ABSTRACT

The invention relates to a process for the fluid catalytic cracking of hydrocarbons using a downflow reactor and a catalyst with an Akzo Accessibility Index (AAI) of at least 3.5. This process combines the advantages of the use of downflow and riser reactors: minimal overcracking and high conversion of the higher boiling fraction. The use of a downflow reactor minimizes overcracking, while the cracking catalyst facilitates high conversions, even for high-boiling fractions in the feed.

CROSS REFERENCE TO RELATED APPLICATIONS

[0001] This application claims priority from EPO Application No.01202203.4, filed Jun. 8, 2001, and PCT Application No. PCT/EPO2/05745,filed May 24, 2002, the contents of which are incorporated by referenceherein.

BACKGROUND OF THE INVENTION

[0002] 1. Field of the Invention

[0003] The present invention relates to a process for fluid catalyticcracking (FCC) of hydrocarbon feeds in a downflow reactor using aspecified cracking catalyst.

[0004] 2. Prior Art

[0005] FCC processes are well known. In the more usual FCC processesemploying riser reactors the catalyst and the hydrocarbon feed flowupward, while in FCC processes employing downflow reactors the catalystand the hydrocarbon feed flow downward.

[0006] In riser reactors solids flow upward due to the lift caused bythe ascending vaporised feed. However, the velocity of the hydrocarbonfeed is lower near the wall than it is near the center of the reactor.Therefore, the catalyst will move more slowly near the reactor wall thannear the center, resulting in a slower moving area with a high catalystdensity near the wall and a low-resistance path of ascending feed nearthe center. Hence, the feed mainly flows through the center, whereas thecatalyst is mainly located near the walls. The resulting flow pattern iscalled core-annulus.

[0007] Furthermore, the upward flow of solid catalyst and hydrocarbonvapor in riser reactors opposes gravity, resulting in a catalyst flowthat is significantly slower than the much lighter hydrocarbon flow. Theratio of feed velocity to catalyst velocity, i.e. the slip factor,generally is about 2-3. This results in backmixing of the catalyst,leading to longer residence times of the catalyst and the occurrence ofundesirable secondary reactions (overcracking).

[0008] In contrast to riser reactors, downflow reactors do not displaylarge differences in velocity and catalyst density between the centerand the wall of the reactor. Furthermore, as the catalyst particles donot oppose gravity, the difference in velocity between the catalyst flowand the hydrocarbon flow in these reactors is smaller than in riserreactors. The slip factor of downflow reactors generally is about 1.

[0009] Consequently, backmixing is largely avoided, the catalyst isdistributed more evenly across the entire reactor, and the effectivecontact time of the catalyst and the feed in a downflow reactor is lessthan in a riser reactor. Although this reduces the formation ofby-products, it also results in a decrease in the conversion of mainlythe larger, higher-boiling compounds.

[0010] For prior art publications dealing with cracking units comprisingdownflow reactors, reference is made to U.S. Pat. No. 5,449,496, U.S.Pat. No. 5,582,712, U.S. Pat. No. 6,099,720, U.S. Pat. No. 5,660,716,U.S. Pat. No. 5,951,850, and EP 0 254 333.

[0011] Of these publications only a few focus on optimizing the processby way of using a specified catalyst. In U.S. Pat. No. 5,660,716 use ismade of a low acidity catalyst. It is recommended to use it inconjunction with high temperatures and high catalyst-to-oil ratios toobtain acceptable conversion levels. A similar teaching—including theuse of catalyst-to-oil ratios of 25 to 80 w/w %—is contained in U.S.Pat. No. 5,951,850, in which it is recommended to use a catalystcontaining a zeolite having a unit cell size of at most 24.50 Angstroms.The high catalyst-to-oil ratio may jeopardize the performance of theunit as far as catalyst separation, stripping, and regenerationcapabilities are concerned. Moreover, wear of the equipment caused bythe catalyst may become critical.

[0012] In sum, the teachings of the prior art tend into the direction ofusing low activity catalysts in conjunction with high temperatures andhigh catalyst-to-oil ratios to compensate for the lower catalystactivity.

SUMMARY OF THE INVENTION

[0013] In one embodiment, the present invention comprises a process forcracking hydrocarbon feeds which combines the advantages of downflow andriser reactors: minimal overcracking and high conversion of thehigher-boiling fraction. The process comprises the following steps:

[0014] a) atomizing and injecting a hydrocarbon feedstock into the topportion of a tubular downflow reactor and contacting this hydrocarbonfeedstock with a catalyst having an AAI of at least 3.5,

[0015] b) separating reaction products and spent catalyst at the bottomof said downflow reactor,

[0016] c) treating the spent catalyst with steam,

[0017] d) regenerating the spent catalyst in a regeneration zone, and

[0018] e) recycling the regenated catalyst to the downflow reactor.

[0019] Other embodiments of the invention involve catalystcharacteristics and composition, and process details, all of which arediscussed in detail hereinbelow.

DETAILED DESCRIPTION OF THE INVENTION

[0020] The use of a downflow reactor minimizes overcracking, while thehigh accessibility of the cracking catalyst facilitates highconversions, even for high-boiling fractions in the hydrocarbon feed.

[0021] The AAI is a measure of the accessibility of the catalyst poresto large, often high-molecular weight compounds and can be determinedaccording to the method described in non pre-published European patentapplication No. 01202147.3 filed on Jun. 5, 2001, which application isincorporated by reference. This method involves adding the porousmaterial to a stirred vessel containing large, preferably rigid, andoften high-molecular weight compounds dissolved in a solvent andperiodically analysing the concentration of these compounds in thesolution. The relative concentration of the large compounds (in %) canbe plotted against the square root of time (in minutes). The AAI isdefined as the initial slope of this plot.

[0022] The higher the AAI value, the more accessible the catalyst poresare.

[0023] If the pores of the cracking catalyst are highly accessible toeven the higher-boiling fractions of the hydrocarbon feed, the feedmolecules will diffuse quickly through the pores and optimum use is madeof the active sites present in the catalyst pores. Hence, highconversions can be reached with such catalysts.

[0024] It is emphasized that the AAI is not equivalent to the porevolume of a catalyst. The AAI deals with the accessibility of this porevolume, e.g. the size of the pore entrance. Hence, catalysts with a highpore volume can have low AAI values if the pore entrances are narrow.

[0025] According to the process according to the invention, ahydrocarbon feedstock is atomized and injected into the top portion of atubular downflow reactor, thereby contacting this hydrocarbon feedstockin the absence of added hydrogen with a hot, fluidized stream ofcatalyst having an AAI of at least 3.5. Next, the spent catalyst, havingcoke and hydrocarbonaceous material deposited thereon, is separated fromthe reaction products. The hydrocarbonaceous material is stripped fromthe spent catalyst by treatment with steam. The coke is removed from thespent catalyst during the regeneration step, involving combustion of thecoke in an oxygen-containing atmosphere at a temperature of about600-850° C., preferably 650-750° C. Finally, the regenerated catalyst isrecycled to the downflow reactor.

[0026] The catalyst-oil contact time preferably is 0.5 to 5 seconds,more preferably 0.5 to 4 seconds, and even more preferably 1 to 3seconds. The temperature at the reactor outlet preferably is between 450and 700° C., more preferably between 500 and 600° C. The catalyst/oilratio preferably is between 2 and 15. The spent catalyst is continuouslyremoved from the reaction zone and made up with catalyst essentiallyfree of coke resulting from the regeneration zone. To make up forcatalyst losses, fresh catalyst is regularly added to the process. Ifdesired, part of the catalyst inventory can be withdrawn and replaced byfresh catalyst to adjust, e.g., the activity, selectivity or metalcontamination of the circulating catalyst inventory.

[0027] The fluidisation of the catalyst with various gas streams allowsthe transport of the catalyst between the reaction zone and theregeneration zone.

[0028] The Catalyst

[0029] The AAI of the catalyst to be used in the process according tothe invention is at least 3.5, preferably at least 5.0, more preferablyat least 6.0. The maximum AAI value depends on the required physicalproperties, such as apparent bulk density and friction strength.

[0030] The catalyst preferably comprises 10-60 wt. % of a solid acid,0-50 wt. % of alumina, 0-40 wt. % of silica, and the balance kaolin.More preferably, the catalyst comprises 20-50 wt. % of solid acid, 5-40wt. % of alumina, 5-25 wt. % of silica, and the balance kaolin. Mostpreferably, the catalyst comprises 25-45 wt. % of solid acid, 10-30 wt.% of alumina, 5-20 wt. % of silica, and the balance kaolin.

[0031] The catalyst may comprise solid acid, matrix, and/or any othercomponent commonly used in FCC catalysts such as metal passivatingagents.

[0032] The matrix typically contains silica, alumina, silica-alumina,and/or clay. A preferred clay is kaolin.

[0033] The solid acid can be a zeolite, e.g., a ZSM-type zeolite such asZSM-5 or a faujasite-type zeolite, a silicoaluminophosphate (SAPO), analuminophosphate (ALPO), or a combination thereof. Preferably, the solidacid is a zeolite, more preferably a faujasite-type zeolite. The zeoliteis optionally ultrastabilised and/or rare earth exchanged, e.g. zeoliteY, zeolite USY, zeolite REY, or zeolite REUSY. The rare earth content ofthe zeolite preferably is below 16 wt %.

[0034] The micropore volume of the catalyst preferably is at least 0.050ml/g, whereas the external surface area preferably is at least 100 m²/g.

[0035] Suitable methods for the preparation of such highly accessiblecatalysts include the methods disclosed in Brazilian patent publicationBR PI 9704925-5A and in non pre-published European patent applicationNo. 01202146.5, filed on Jun. 5, 2001, which applications are bothincorporated by reference.

[0036] The first method comprises mixing the catalyst components orprecursors thereof in an aqueous slurry to form a precursor mixture,adding a pore-forming agent to this mixture, followed by shaping, e.g.spray-drying.

[0037] The pore-forming agent controls the porosity of the catalyst. Apreferred pore-forming agent is a water-soluble carbohydrate, e.g.,sucrose, maltose, cellobiose, lactose, glucose, or fructose. Thesepore-forming agents can be readily removed after the catalystpreparation. Thermogravimetric analyses indicate that the pore-formingagent can be removed to less than 5 wt. % remaining in the catalyst.

[0038] According to the second method, the catalyst components orprecursors thereof are mixed in an aqueous slurry to form a precursormixture, the mixture is fed to a shaping apparatus and shaped to formparticles, in which process just before being fed to the shapingapparatus the mixture is destabilized, i.e. its viscosity is increased.

[0039] More in particular, this method involves feeding suspendedcatalyst components or precursors thereof from one or more vessels (the“holding vessels”) via a so-called pre-reactor to a shaping apparatus.In this pre-reactor the catalyst precursor mixture is destabilized.

[0040] In this specification a destabilized mixture is defined as amixture whose viscosity is higher after leaving the pre-reactor (andbefore shaping) than before entering the pre-reactor. The viscosityincrease is due to induced polymerisation or gelling of catalyst bindermaterial in the pre-reactor. The viscosity is typically increased from alevel of about 1-100 Pa·s at a shear rate of 0.1 s⁻¹ before entering thepre-reactor to a level of about 50-1,000 Pa·s or higher at a shear rateof 0.1 s⁻¹ after leaving the pre-reactor. In any case, it is preferredto induce a viscosity increase of at least 10 Pa·s, more preferably atleast 50 Pa.s, and most preferably at least 100 Pa·s (measured at ashear rate of 0.1 s⁻¹). Preferably, the viscosity is increased from alevel of about 1-50 Pa·s at a shear rate of 0.1 s⁻¹ before entering thepre-reactor to a level of about 50-500 Pa·s at a shear rate of 0.1 s⁻¹after leaving the pre-reactor. The viscosity can be measured by standardrheometers, such as plate-and-plate rheometers, cone-and-platerheometers or bop-and-cup rheometers.

[0041] Destabilization of the catalyst precursor mixture is performed inthe pre-reactor just before feeding the mixture to the shapingapparatus. The time period involved, i.e. the time which elapses betweenthe start of the destabilization and the shaping, depends on the exactconfiguration of the pre-reactor and on the time needed thereafter forthe destabilized mixture to reach the shaping apparatus. Time periods ofup to half an hour are possible, but may be less preferred foreconomical reasons. Preferred is a time period of less than 300 seconds.A more preferred time period is less than 180 seconds. Destabilizationcan be performed by temperature increase, pH increase or pH decrease,and/or the addition of gel-inducing agents such as salts, phosphates,sulphates, and (partially) gelled silica.

[0042] A suitable shaping method is spray-drying. For more detailsconcerning this method we refer to non pre-published European patentapplication No. 01202146.5.

EXAMPLES Comparative Example 1

[0043] This Example compares the performance of a conventional catalystin a downflow and a riser reactor.

[0044] A conventional equilibrium catalyst was evaluated in two distinctpilot units, one comprising a downflow reactor and the other comprisinga conventional riser reactor. Both units operated at the same reactiontemperature. The properties of the gas oil used are listed in Table 1.TABLE 1 Physical and chemical properties of the gas oil used API 18.6Density 20/4° C. (g/ml) 0.9386 Viscosity (ASTM D445) (cSt) 268 Anilinepoint (° C.) 80.8 Basic Nitrogen (ppm) 961 Concarbon Residue (wt %) 0.38Initial Boiling Point, IBP (° C.) 309 Final Boiling Point, FBP (° C.)602

[0045] Table 2 displays the results of the cracking process using theriser and the downflow reactor at constant coke production. From theseresults it follows that the use of a downflow reactor leads to improvedconversion levels and improved selectivity to C₃ olefins, as well as toimproved hydrogen selectivity. However, the bottoms conversion in thedownflow reactor is lower than in the riser reactor. TABLE 2 Riserreactor Downflow reactor Reaction Temperature (° C.) 550 550 Conversion(wt %) 72.6 74.8 Catalyst/Oil Ratio, CTO (wt/wt) 7.8 8.7 Delta Coke(Coke/CTO, wt %) 1.13 1.02 Coke (wt %) 8.8 8.8 Fuel Gas (wt %) 4.8 4.8Hydrogen (wt %) 0.60 0.17 LPG (wt %) 17.9 20.4 Propene (wt %) 4.84 6.63Gasoline (wt %) 41.0 40.8 LCO (wt %) 15.9 12.3 Bottoms (wt %) 11.6 13.0

Comparative Example 2

[0046] A catalyst was prepared in the following way:

[0047] A silica hydrosol was prepared by the controlled neutralisationunder acidic pH of a sodium silicate solution by diluted sulfuric acid.To the freshly prepared hydrosol were added, sequentially and underthorough agitation, powdered kaolin, an acidic suspension of aboehmite-type alumina, and an acidic suspension of REY-zeolite. Theresulting precursor suspension had a solids content of 20 wt %.

[0048] The precursor mixture was subsequently fed to a spray-dryer andcatalyst microspheres were recovered. The microspheres were re-suspendedin ammoniated water and filtered under reduced pressure. The so-formedfilter cake was twice exchanged with an ammonium sulfate solution at 45°C. and washed three times with water at the same temperature. Finally,the catalyst particles were dried in an oven under circulating air at110° C. for 16 hours, which yielded the fresh sample EC2.

[0049] EC2 was composed of 40 wt % of ultrastabilised Y-zeolite with aSAR of 5.5 and exchanged to reach 5 wt % rare earth oxides (RE₂O₃); 30wt % silica-alumina matrix; and 30 wt % kaolin.

[0050] The physical properties of this catalyst are displayed in Table3.

EXAMPLE 1

[0051] The catalyst of this Example was prepared using exactly the sameprocedure as that of Comparative Example 2, except that—as taught inBrazilian PI BR 9704925-5A—sucrose was added to the precursor mixture.This resulted in catalyst E1. The physical properties of this catalystare displayed in Table 3. TABLE 3 BET MiPV MSA ABD (m²/g) (ml/g) (m²/g)(kg/dm³) AAI EC2 287 0.103  66 0.71 2.0 E1 362 0.115 110 0.70 6.0

[0052] In this Table, BET is the well-known BET surface area, MiPV isthe micropore volume, and MSA the mesopore (20-500 Å) surface area, alldetermined by N₂ adsorption (t-plot method).

[0053] ABD stands for the Apparent Bulk Density, which is defined as themass of catalyst per unit of volume in a non-compacted bed. The ABD ismeasured after filling a gauged cylinder of fixed, pre-determined volumewithout compaction of the bed.

[0054] The AAI was determined by preparing a 1 l solution of 15 g KuwaitVGO in toluene by heating a Kuwait VGO feed to 70° C. in an oven. 15 gof the warm Kuwait VGO were suspended in 200 ml warm toluene. Themixture was well stirred and adjusted to 1 litre with toluene. Thesolution was stored in the dark. 50.00 g of this solution were added toa 100 ml beaker (glass) connected to a peristaltic pump and a detectorby way of tubes. The solution was stirred with a propeller stirrer at400 rpm and the peristaltic pump was set at 21 g/min. Aspectrophotometer was used as detector. This spectrophotometer was setto zero using a toluene solution.

[0055] Next, 1 g of a 53-75 microns sieve fraction of the catalyst wasadded to the Kuwait VGO in toluene solution. Once per second theasphaltene concentration was measured by spectrophotometry at awavelength of 560 nm. After 5 minutes, the measurement was stopped andthe relative absorbance was plotted versus the square root of time. Theslope, i.e. the Akzo Accessibility Index (AAI), was determined.

EXAMPLE 2

[0056] Portions of catalysts EC2 and E1 were hydrothermally deactivatedusing a 100% steam atmosphere at 788° C. for 5 hours in order tosimulate the equilibrium state. The resulting deactivated catalysts arecalled EC2D and E1D, respectively.

[0057] The deactivated samples were tested in the same downflowreactor-containing unit and under the same conditions as in ComparativeExample 1. The results at the same conversion levels and the same cokelevels are listed in Table 4.

[0058] The unit inventory was 2 kg and the gas oil flow rate was 1.7kg/h. The operating conditions were: reaction pressure 0.1 kgf/cm²g,contact time 2 seconds, temperature at the reactor exit 540° C. and inthe stripper 500° C. The catalyst/oil ratio (wt/wt) was varied in therange 6-9 by altering the feed temperature in the adiabatic reactor.TABLE 4 EC2D E1D Equal Conversion (wt %) 77.0 77.0 Catalyst/Oil Ratio(wt/wt) 8.6 6.0 Coke (wt %) 8.9 7.8 Fuel Gas (wt %) 4.0 3.2 Hydrogen (wt%) 0.10 0.04 LPG (wt %) 19.0 17.4 Propene (wt %) 5.1 4.5 Gasoline (wt %)45.1 48.6 LCO (wt %) 12.2 12.6 Bottoms (wt %) 10.8 10.4 Equal Coke (wt%) 8.0 8.0 Catalyst/Oil Ratio (wt/wt) 6.4 6.9 Conversion (wt %) 72.279.7 Fuel Gas (wt %) 3.6 3.3 Hydrogen(wt %) 0.10 0.04 LPG (wt %) 17.019.3 Propene (wt %) 4.4 4.9 Gasoline (wt %) 43.6 49.0 LCO (wt %) 13.212.0 Bottoms (wt %) 14.7 8.4

[0059] From this Table it is clear that a process using a combination ofa downflow reactor and a catalyst with an AAI of at least 3.5 results inhigh conversion levels and gasoline yields, combined with high bottomsconversion. These results were obtained using reaction temperatures andcatalyst-to-oil ratios which are normally employed in industrial riserreactors. Moreover, the feed used was one having a high basic nitrogencontent.

[0060] On using the catalyst of Example 1 it is possible to operate in adownward flow reactor without conversion loss at lower catalyst-to-oilratios than those recommended in the prior art.

[0061] The results also show a tendency towards improvements in coke,fuel gas and gasoline selectivity.

[0062] At constant coke the synergism between the use of a downflowreactor and a catalyst having an AAI of at least 3.5 is especiallybeneficial.

[0063] Fuel gas and hydrogen yields are reduced and light olefins areincreased. Compared to the base case bottoms conversion is increasedtoo. This shows that the disadvantage in respect of bottoms conversionfor downflow operations observed in the base case—see Table 2—may befully compensated by the process according to the invention.

[0064] Finally, it has been observed that the strippability of catalystshaving an AAI of at least 3.5 is greatly improved in comparison withprior art catalysts not having such high accessibility.

1. A process for the fluid catalytic cracking of hydrocarbons comprisingthe following steps: a) atomizing and injecting a hydrocarbon feedstockinto the top portion of a tubular downflow reactor and contacting thishydrocarbon feedstock with a catalyst having an AAI of at least 3.5, b)separating reaction products and spent catalyst at the bottom of saiddownflow reactor, c) treating the spent catalyst with steam, d)regenerating the spent catalyst in a regeneration zone, and e) recyclingthe regenated catalyst to the downflow reactor.
 2. The process of claim1 wherein the catalyst has an AAI of at least 5.0.
 3. The process ofclaim 2 wherein the catalyst has an AAI of at least 6.0.
 4. The processof claim 1 wherein the catalyst has been obtained by combining catalystcomponents or precursors thereof in an aqueous medium to form a catalystprecursor mixture, feeding the mixture to a shaping apparatus, andshaping the mixture to form particles, in which process just beforebeing fed to the shaping apparatus the mixture is destabilized.
 5. Theprocess of claim 1 wherein the catalyst has been obtained by mixing thecatalyst components or precursors thereof in an aqueous mixture, addinga pore-forming agent to this mixture, followed by shaping.
 6. Theprocess of claim 5 wherein the pore-forming agent is a water-solublecarbohydrate.
 7. The process of claim 1 wherein the catalyst comprises10-60 wt. % of a solid acid, 0-50 wt. % of alumina, 0-40 wt. % ofsilica, and the balance kaolin.
 8. The process of claim 7 wherein thecatalyst comprises 20-50 wt. % of solid acid, 5-40 wt. % of alumina,5-25 wt. % of silica, and the balance kaolin.
 9. The process of claim 8wherein the catalyst comprises 25-45 wt. % of solid acid, 10-30 wt. % ofalumina, 5-20 wt. % of silica, and the balance kaolin.
 10. The processof claim 7 wherein the solid acid is selected from the group consistingof ZSM-type zeolites, faujasite-type zeolites, silicoaluminophosphate(SAPO), aluminophosphate (ALPO), and combinations thereof.
 11. Theprocess of claim 10 wherein the solid acid is a rare earth exchangedzeolite.
 12. The process of claim 11 wherein the rare earth content ofthe zeolite is below 16 wt %.